process stimulation

PRE-FEED YAMAMA FORMATION DEVELOPMENT

WEST QURNA 2 PROJECT, IRAQ

Process Simulation Report

A 18-10-2012 IDCInter Discipline

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8015 0152 WPAD 00 000 PC RP 00004 EN

8015-0152-WPAD-00-000-PC-RP-00004 Process Simulation Report Rev A.docx

LUKOIL MID-EAST LIMITED

PRE-FEED SERVICES CONTRACTYAMAMA FORMATION DEVELOPMENT


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Doc. title Process Simulation Report Rev.: A Status: IDCCOMPANY No.:Contractor No.:

8015-0152-WPAD-00-000-PC-RP-00004N/A Page: of 28

DOCUMENT REVISION HISTORY SHEET

ИСТОРИЯ ИЗМЕНЕНИЙ

REV STATUS DATE ISSUED UPDATE / AMENDMENT DETAILS

РЕВ СТАТУС ДАТА ВЫПУСКАИНФОРМАЦИЯ ОБ ИЗМЕНЕНИЯХ/ ПОПРАВКАХ

A IDC 18-10-2012 ISSUED FOR INTER DISCIPLINE CHECK





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CONTENT

1 INTRODUCTION....................................................................................................4

1.1 Purpose.............................................................................................................................................................4

1.2 Scope.................................................................................................................................................................4

1.3 Definitions and abbreviations.........................................................................................................................5

2 REFERENCE DOCUMENTS.................................................................................6

3 SIMULATION BASIS.............................................................................................7

3.1 SOFTWARE and Model.................................................................................................................................7

3.2 Formation Fluid Definition.............................................................................................................................7

3.3 Equipment Design............................................................................................................................................73.3.1 Compressor...................................................................................................................................................73.3.2 Heat Exchanger.............................................................................................................................................7

3.4 Product Specification & Battery Limit Conditions......................................................................................7

3.5 List of Simulation Case...................................................................................................................................7

3.6 Sensitivity Study..............................................................................................................................................7

4 RESULT AND DISCUSSION.................................................................................8

4.1 Oil Separation Unit..........................................................................................................................................84.1.1 Operating pressure of HP Seperator..............................................................................................................84.1.2 Operating pressure of Stabilizer....................................................................................................................8

4.2 Gas Compressor Unit......................................................................................................................................84.2.1 Interstage Liquid Routing.............................................................................................................................8

4.3 Gas Sweetening Unit........................................................................................................................................8

4.4 Gas Dehyration Unit........................................................................................................................................8

4.5 HC Dew Point Control Unit............................................................................................................................8

4.6 Deethaniser Column........................................................................................................................................8

4.7 Depropaniser Column.....................................................................................................................................8

4.8 Debutaniser Column........................................................................................................................................8

4.9 Propane Regfrigeration Unit..........................................................................................................................8

4.10 Sensitivity Study with PVT Data....................................................................................................................8





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1 INTRODUCTION

LukOil Middle East Limited (LME), in partnership with the South Oil Company of Iraq, is currently planning to develop its upstream production facilities and infrastructure to meet a target of 1.8MMBOPD of production in the West Qurna 2 oil field in Southern Iraq. The field consists of two primary reservoirs, Mishrif and Yamama, which will be developed as separate projects. It is envisaged that the Yamama reservoir will produce up to 1.1MMBOPD of the 1.8MMBOPD targeted from the entire field. Presently the Yamama development is in the Concept phase of project execution and design. Construction of the preliminary production facilities for the Mishrif reservoir is currently underway.

WorleyParsons has been contracted to execute the Pre-FEED or SELECT phase of the Yamama Development Project. The main focus of this phase will be to frame and assess each of the concept development cases for the future development of the Yamama reservoir. The development cases shall include evaluation of options for wells, gathering systems, gas and water reinjection, gas and water treatment, expansion of the Mishrif power plant, a light oil export pipeline and tank farm upgrades, tie-ins to the natural gas liquids and gas export systems and further expansion of supporting facilities and infrastructure as necessary.

The Project aims to produce a detailed Decision Support Package identifying the preferred design options, and comprehensive FEED ITB documents for progression of the design to the FEED phase.

1.1 Purpose

This document presents the basis of steady state process simulation for YAMAMA facility which will be used for generation of Heat & Material Balance, design of equipment and utilities calculation. This will be used as basis by the FEED contractor for a more detailed study and optimization during execution of the

FEED.

1.2 Scope

The scope of this document includes simulation comprising of the following facilities

Central Processing Facility Gas Treatment Plant





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1.3 Definitions and abbreviations

The following definitions and abbreviations are used in this document:

BOD Basis of Design

CPF Central Processing Facility

FEED Front End Engineering Design

GOR Gas to Oil Ratio

GTP Gas Treatment Plant

MMBOPD Million Barrel Oil Per Day

MMSCFD Million Standard Cubic Feet Per Day

LME Lukoil Middle East Limited

LPG Liquefied Petroleum Gas

TOR Terms of Reference

WP Worley Parsons





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2 REFERENCE DOCUMENTS

2.1 PROJECT REFERENCE DOCUMENT

1. Framing Terms of Reference (Doc No: 8015-0152-WPAD-00-000-PC-RP-00001)2. Preliminary Basis of Design (Doc No: 8015-0152-LMEL-00-000-PC-BD-00001)

2.2 PRACTICES, CODES AND STANDARDS

2.3 Order of Precedence

The precedence applying for use of the Codes, Standards, Specification and Statutory requirements for this project is as follows:

Statutory Requirements, Local applicable laws and Regulation in Iraq Iraq Standards Project Specifications and Standards Project Datasheets International Standards Service Authority Standards

In the event of an inconsistency, conflict or discrepancy between any of the Standards, Specifications and Statutory requirements, the most stringent and safest requirement applicable to the project will prevail to the extent of the inconsistency, conflict or discrepancy. Any inconsistencies, critical to the design, shall be brought to the attention of LME for resolution.


Associated Gas

Crude

Gas Formation Fluid (Dry Basis)

Flashed Liquid

Flashed Gas

Water

Formation Fluid (Wet Basis)

Flashed @ Std. Condition

Adj-1

Adj-2

Adj-3

Choke valve

Oil Separation train

Produced water

Adj-4

Stabilized Crude




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3 SIMULATION BASIS

3.1 SOFTWARE and Model

HYSYS version V7.3 is used for the simulation of Central processing Facility and Gas Treating Plant. A complete simulation flow sheet for a single processing train is prepared covering all the units as sub-flowsheet in HYSYS. Peng-Robinson (PR) Thermodynamic Fluid Property Package is used for all units except for Gas sweetening and Dehydration units. The Peng-Robinson (PR) model is ideal for VLE calculations for hydrocarbon systems. The PR model rigorously solves any single, two or three phase system with a high degree of efficiency and reliability and is applicable over a wide range of conditions. The applicable range for Peng Robinson equation is given as below:

Temperature Range > -271°C or -456°F

Pressure Range < 100,000 kPa or 15,000 psia

Gas Sweetening unit is better modeled using specialized software like PROMAX, however to maintain the continuity of the simulation flow sheet, Gas sweetening is modeled in HYSYS with Amine package. However, the HYSYS results are compared with PROMAX and suitable modifications (ex. correcting amine recirculation flow rates, amine mix proportions etc.) are incorporated in HYSYS model to represent the reality.

For Gas Dehydration unit, HYSYS Glycol package is used.

3.2 Formation Fluid Definition

The characterization of formation fluid composition is very critical as it forms the basis for the complete simulation. The steps followed to define the formation fluid composition are outlined below with schematics.





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STEPS:

1. Characterize the crude stream based on Assay Data provided in BOD (Table 6-1)

2. Model Gas stream based on Average Composition (Layer A+B) provided in BOD (Table 6-4)

3. Mix the above crude and gas streams and flash at standard conditions

4. Adjust the crude flow rate (Adj-1) to match the required stabilized crude flow rate from Oil separation train (i.e., one train crude capacity)

5. Adjust Gas flow rate (Adj-2) to match the required GOR based on Flashed Gas & Liquid

6. Mix the above adjusted crude and gas streams to get the formation fluid at the design oil rate and GOR on dry basis.

7. Mix the formation fluid (dry basis) with water to obtain the formation Fluid on Wet Basis. Adjust the water flow rate (Adj-3) to match the required water cut in the formation fluid (wet basis).

8. Adjust Temperature and Pressure (Adj-4) of formation fluid (wet basis) to match well free flowing conditions given in BOD (Table 7-2).

3.3 Equipment Design

The equipment design basis which will have major impact on the simulation flow sheet is discussed as follows:

3.3.1 Compressor

The number of stages of compressing is calculated to restrict the discharge temperature in each stage to less than 180°C with the consideration of inter-stage cooling facility.

3.3.2 Heat Exchanger

3.3.2.1 Air Cooled Exchangers:

As per table 5.3.1 given in BOD, the ambient air temperature for Air cooler design is 46°C. Based on a typical approach temperature of 10°C, the process side outlet temperature for Air coolers is considered as 56°C.

3.3.2.2 Cooling Water Exchangers:

Assuming a wet bulb temperature of 30°C and maximum approach temperature in the cooling tower as 3°C, the cooling water supply temperature to heat exchangers is considered as 33°C. A maximum





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temperature rise of 10°C is considered in a single user in view of cooling tower design and corrosion rate in the cooling water system.

3.4 Pressure profile in the simulation flow sheet

The basis for operating pressure in different sections of the units is discussed in detail in respective sections however in general, the following basis is considered:

1) Operating pressure of the separation column is determined by the cooling medium temperature available for providing the condenser duty and dis-integration temperature of the column bottom product.

2) Units which involve absorption such as Dehydration & Gas Sweetening are favored by high operating pressure and units involving stripping (Stabilization) are better operated at low pressure. Accordingly, the pressure levels are optimized.

3) A maximum allowable pressure drop of 1 bar is considered in equipment such as heat exchangers (except condensers and reboilers), filters etc. and interconnecting piping between units to calculate the pressure profile in the flow sheet.

3.5 Simulation Cases & Basis

3.5.1 Water Injection Concept:

As per the client response to TQ no: 00-TQ-WPAD-LMEL-AL-0014, CPF train size is standardized to 130 MBOPD. Accordingly, there will be nine (9) numbers of CPF trains in the water injection concepts to meet the design capacity of 1.1 MMBOPD with 95% availability factor.

With respect to the GTP, three (3) numbers of trains are considered, such that 3 numbers of CPF trains are served by 1 train of GTP. In the simulation flow sheet, one train of CPF and one train of GTP are modeled thus; the stabilized crude produced from CPF is 130 MBOPD and sales gas flow rate from GTP is 213 MMSCFD. This case is referred to as WIS Simulation Case-1.

Additionally to cover the TOR case related to consideration of 1 st stage separation at the Well Pads, the WIS simulation case-1 is modified accordingly and referred as WIS Simulation Case-2.

A constant GOR ratio of 135 sm3/sm3 is considered for Water Injection concept design as per the client response to TQ no: 00-TQ-WPAD-LMEL-AL-0006.

3.5.2 Gas Injection Concept:

Similar to Water Injection Concept, nine (9) numbers of CPF trains are considered in Gas Injection concept. The associated gas will be compressed and Dehydrated in the CPF and routed to Gas Injection Stations. As the GOR ratio in the Gas Injection concept varies significantly from 136 sm3/sm3 (Year 2017) to 385 sm3/sm3 (Year 2034), phasing concept is applied for associated gas compressor, dehydration units and Gas Injection Stations. As per the analysis of GOR profile (refer section x for details), the associated gas compressor and Dehydration capacity per train is considered as 326 MMSCFD and individual Injection Gas compressor capacity is considered as 256 MMSCFD. Accordingly,





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there will be six (6) numbers of associated gas compressor & Dehydration trains and fourteen (14) numbers of Injection Gas compressors.

In the simulation flow sheet, one train of CPF including one train of associated gas compressor, Dehydration unit and one gas injection station compressor are modeled. This case is referred to as GIS Simulation Case-1.

Additionally to cover the TOR case related to consideration of 1 st stage separation at the Well Pads, the GIS simulation case-1 is modified accordingly and referred as GIS Simulation Case-2.





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4 RESULT AND DISCUSSION

4.1 Water Injection Concept Simulation (WIS Simulation Case-1)

The discussion section is structured to cover the results of simulation study following the flow scheme starting from 1st stage separation. The optimization related to individual system is discussed in the respective section.

4.1.1 Oil Separation Unit

The objective of the Oil separation train simulation is to separate the associated gas and water from the crude to meet the crude specification mentioned in the below table as per BOD (Table 8-2). To achieve H2S and RVP specifications, the crude needs to be flashed in series of separators and treated in Desalter to meet water, salt and BS&W specification. The optimizations related to stages of separation and operating pressure of the stages are discussed in the following sections.

Parameter Unit Specification

H2S ppmw <50

RVP of light oil & regular oil (RVP@38ºC) kPa 42.8 to 44.2

Salt Content mg/L <28.5

Water Content Vol % <0.15

Bottom Sediments and Water [BS&W] Vol % <0.5

4.1.1.1 1st Stage Separation

Operating Pressure:

Higher operating pressure in the 1st stage separator will require a smaller size for the gathering network however it will have impact on other parameters such as Gas compressor load, gathering network design pressure, wall thickness etc. Hence, to optimize the 1 st stage separator operating pressure, the following basis is considered and the aforesaid parameters are tabulated to derive conclusion.

Basis:

Three stage separations considered with 2nd stage (LP separator) and 3rd stage (Stabilizer column) operated at 5.5 bara in all options. Only the 1st stage Separator operating pressure is considered at four different levels as described below:

Option-1: 1st Stage separator at 41 bara

Option-2: 1st Stage separator at 32 bara (limiting to max limit of 300 # rating in the gathering network)

Option-3: 1st Stage separator at 15.5 bara





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Option-4: 1st Stage separator at 5.5 bara

The HSYSY results** are tabulated in the below table:

Parameters

Option 1

(1st Stage – 41 bara)

Option 2

(1st Stage – 32 bara)

Option 3

(1st Stage – 15.5 bara)

Option 4

(1st Stage – 5.5 bara)

Choke Downstream Pressure, barg

53 44 27.5 17.5

Gathering Line Rating 600 # 300 # 300 # 150 #

Gathering Line Size Smallest Small Big Bigger

1st Stage Separator Rating 300 # 300 # 300 # 150 #

1st Stage Separator Size (ID x L), m 4.3 x 13.1 4.3 x 13.1 4.4 x 13.2 4.5 x 14.2

2nd Stage Separator Pressure, bara 5.5 5.5 5.5 -

Stabilizer Pressure, bara 5.5 5.5 5.5 5.5

Stabilizer Size (ID x H), m 5.3 x 18.7 5.2 x 19 5.2 x 19 5.0 x 19.6

Design Pressure & Wall thickness Highest Higher Low Lower

Associated Gas Compr. Load, MW 70 69 63 118

Expected CAPEX Higher Higher Lower Lower

Expected OPEX Higher Higher Lowest Highest

** The equipment sizes and load values specified in the above table are for one train. Train size was not standardized during the optimization study and hence the values specified in the table may not match exactly with the final simulation case (CPF: 130 MBOPD & GTP: 213 MMSCFD) however for the purpose of optimization the above results are valid. This is applicable for all the optimization studies presented in this report.

It is noticed that with lower operating pressure, the line sizes will increase however the wall thickness will be reduced. Thus, the CAPEX may remain similar for all the cases or it is expected to be higher for increased pressure considering that the wall thickness is governing the CAPEX as assumed in the table above. With this basis, low operating pressures (Option-3 and Optio-4) are considered better than the other two options.

With respect to the OPEX, it is observed that at lowest 1 st stage operating pressure (5.5 bara); the total gas is generated at 5.5 bara and need to be compressed in associated gas compressor from this low pressure. This results in high associated gas compressor duty of 118 MW. With increase in 1 st stage operating pressure up to 15.5 bara, a considerable quantity of associated gas is generated at higher pressure resulting in reduced associated gas compressor duty (63 MW). However, with further increase in operating pressure (> 15.5 bara), the quantity of flashed vapor generated in the 1 st stage separator is very less and most of the gases is again generated at 2nd stage separated operated at 5.5 bara resulting in an upswing in the associated compressor duty (70 MW for 41 bara). Thus, for the design well fluid composition, a 1st stage operating pressure of 15.5 bara generates optimal quantity of associated gas at higher pressure which has the lowest associated compressor duty (63 MW).





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It may be noted that as per BOD (Table 7-2), the Well Tubing pressure for both Water and Gas Injection Station is 200 barg and Choke valve pressure drop is 172 bar. As a result, the pressure at the downstream of the choke valve will be 28 barg. Thus, with the HP separator operating pressure at 15.5 bara (14.5 barg), a reasonable pressure drop of about 14 bar is available in the gathering network (15 to 20 km length).

Hence, based on CAPEX and OPEX discussed above, Option-3 (1st stage separator at 15.5 bara) is the preferred option.

Operating Temperature:

With respect to the temperature, as per BOD (Table 7-2), the well tubing is at 60°C and the calculated choke valve downstream temperature is 61.5°C. Hydraulic simulation with buried pipeline indicates a temperature drop of about 2°C in the between well pad and CPF; accordingly, the HP separator inlet temperature is considered as 58°C.

The selected operating conditions for 1st stage separation is provided below

Parameters 1st Stage Separation

Operating pressure, bara 15.5

Operating Temperature, °C 58

Three phase separation is considered in the 1st stage separator to remove associated gas, oil and produced water. HYSYS will predict ideal separation of water in the separator, hence to represent the reality, a typical concentration of 10 wt% of water is considered in outlet oil from the separator (this will be verified with vendor).

Slug Catcher Requirement:

It may be noted that based on the slug volume estimated by gathering network transient study, a suitable design from the following two options will be considered:

Option-1: 1st Stage separator to be designed to handle slug (if the slug volume is nominal)

Option-2: Inlet Slug Catcher (upstream of 1st stage separator), if the slug volume is high.

Should the Slug catcher be added, the operating pressure can be considered as 16.5 bara, which is 1.0 bar more than the 1st stage separator.

Water Cut Variation:

The water cut in the formation fluid varies from zero (0) to 70 vol% during the plateau oil production period i.e., Year 2021 to 2028. Designing the 1 st Stage separator for 70 vol% will result in an over-sized separator posing under-utilization and low turndown issues during the initial years of formation. Thus, two separators designed to handle the overall water cut range is considered which will be phased in two (2) stages i.e., first Separator will handle the water cut (35 vol%) till Year 2024 and beyond which the





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second Separator will be added to function in parallel with the first separator. The well fluid will be equally distributed between the two separators.

It may be noted that the consideration of two stages of phasing and the rallying period at 2024 (35 vol%) gives the advantage of similar total volumetric flows per separator (before and after Year 2024) resulting in same sizes for both the Separators. Also, the calculated separator size is reasonable (ID: 4.4 m and T/T Length: 13.2m) providing an optimized foot print on the overall plot plan.

2017

2018

2019

2020

2021

2022

2023

2024

2025

2026

2027

2028

2029

2030

2031

2032

2033

2034-10.0

0.010.020.030.040.050.060.070.080.090.0

0

200

400

600

800

1000

1200

Water Cut Variation

Water Cut (vol%) Oil rate (MBOPD)

Wat

er C

ut (v

ol %

)

OIl

rate

(MBO

PD)

4.1.1.2 2st Stage Separation

Three phase separation is considered in the 2nd stage separator so that the water is further removed before the Desalting unit. The flashed gas from the 1st and 2nd stage separators will be routed to the suction of 1st and 2nd stage of associated gas compressors corresponding to the pressure levels. Thus, with the consideration of typical compressor ratio of 3:1 between the 2 stages of compressors and 1 st

stage separator operating pressure optimized at 15.5 bara, the 2nd stage separator operator pressure is fixed at 5.5 bara.

Parameters 2nd Stage Separation

Operating pressure, bara 5.5


As mentioned above, HYSYS will predict ideal separation of water in the separator, hence to represent the reality, a typical concentration of 1 wt% of water is considered in outlet oil from the separator which will be verified with the separator vendor later.

4.1.1.3 Desalter


Second separator to be added in Year 2024




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The objective of the Desalter is to process the oil from 2nd Stage separator to meet the Water, Salt and BS&W specification in the crude. For effective operation, the Desalter needs to be liquid filled without any vapor space at its operating temperature. Considering a typical operating temperature of 120°C, the bubble pressure of the Desalter feed is calculated as 13.7 bara. Allowing a margin of 2 bara, the Desalter operating pressure is fixed at 16.0 bara.

Parameters Desalter

Operating pressure, bara 16


Heat integration consideration to obtain Desalter feed temperature of 120°C is discussed in the subsequent section.

4.1.1.4 3rd Stage Separation

The consideration of number of separation stages and the type of separation (separator versus Trayed column) is critical to meet the H2S and RVP specification in the stabilized crude. The following two options are considered for this optimization.

Option-1: 2 Stages of separators followed by Stabilizer column

Option-2: 3 Stages of separators

Basis:

In both options, 1st stage and 2nd stage separator pressure are considered at 15.5 and 5.5 bara respectively. With respect to the 3rd stage separation, a trayed stabilizer column operating at 5.5 bara is considered in Option-1 and a flash separator operating at 2 bara is considered in Option-2. Refer to Annexure-1 for the Schematics representation of Option-1 & 2.

The HYSYS results are tabulated in the below table:

Parameters2 Stage Separation +

Stabilizer3 Stage Separation (No

Stabilizer)

Hot Oil Duty (CPF), MW 32 84

Compressor Duty, MW 16 30

Compressor Outlet Cooler Duty, MW 33 132

Product Crude RVP, kPaa (42.8 to 44.2) 44 44

Product Crude H2S, ppmw (<50) 18 38

Expected OPEX Lower Higher

Expected CAPEX Lower Higher





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It is noticed that in Option-2 with 3 separators in series, the flashed gas generated by only pressure letdown is very less and consequently H2S and RVP specification in the stabilized crude could not be met. Hence, in addition to pressure letdown, the individual separator inlet stream is heated with hot oil to strip off sufficient vapors to meet the crude specifications. However, this option is found unfavorable as compared to Stabilizer option for the following reasons:

1. The additional heat exchanger at the inlet of each separation stage has increased the inlet temperature and actual volumetric rate to the compressors in Option-2. Also, an additionally compressor is required in option-2 to compress the gas from 3 rd stage separator operating at 2.0 bara. As a result, the overall compressor and its discharge cooler duties in Option-2 are calculated as 162 MW (30 +132 MW) which is significantly higher than 49 MW (16 + 33 MW) obtained in Option-1.

2. The heat input supplied as reboiler duty in a trayed stabilizer column (Option-1) is very effective for achieving the desired separation as compared to feed heating at individual separator inlet in Option-2. As a result, the overall Hot Oil duty requirement in Option-2 (84 MW) is substantially higher as compared to Option-1 (32 MW).

3. The H2S concentration obtained with Stabilizer column (Option-1) provides a considerable cushion to handle any H2S spike in the feed composition in the later years of formation.

Thus, based on the OPEX and CAPEX indications, 2 Stages of separators followed by Stabilizer column (Option-1) is selected.

4.1.1.5 Operating pressure of Stabilizer

The objective of the Stabilizer column is to process the oil form Desalter to strip-off the lighters including the H2S slipped from the upstream flash separators to meet crude product specification. As the basic process is stripping, lower operating pressure favors the column performance and requires lesser reboiler duty (lower bottom temperature) compared to higher operating pressure. However, low operating pressure will have impact opposing effect on the column size (high volumetric rate), compressor load, heat integration etc. Hence, to optimize the Stabilizer operating pressure, the following basis is considered and the aforesaid parameters are tabulated to derive conclusion.

Basis:

For optimization, maximum operating pressure of 5.5 bara has been considered such that the stabilizer overhead gas can be routed to the 1st stage compressor inlet. Additionally two more options with lower operating pressures as mentioned below are considered.

Option-1: Stabilizer operating at 5.5 bara



Refer to Annexure-2 for the Schematics representation of different Options.





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Parameter Option 1 Option 2 Option 3

Stabilizer Pressure, bara 5.5 2.7 1.5

Stabilizer Dimension (ID x H), m 5.2 x 19 3.6 x 26.6 3.2 x 30

Stabilizer bottom Temperature, °C 165 123 100

Stabilizer Reboiler Duty, MW 24.7 13.8 9.8

Additional Exch. Hot Oil Duty, MW - 11.1 30

CPF Hot Oil Duty, MW 24.7 24.9 39.8

GTP Feed Gas Comp. Duty, MW 8 9.7 14.2

Total Energy (Hot Oil + GTP Feed Gas Compressor), MW

32.7 34.6 54

Stabilized Crude Air Cooler Duty, MW 7 7 7

Stabilized Crude Cooling Water Duty, MW 12 12 12

S&T Exchanger Area, m2 9,500 5,300 2,900

Air cooler area, m2 / Fan power kW 17,000 / 150 17,000 / 150 17,000 / 150

Crude RVP, kPaa (42.8 to 44.2) 44 44 25.6

Product Crude H2S, ppmw (<50) 17.5 49.9 49.9

Additional Equipment Stabilizer Feed Bottom

Exchanger

Desalter Feed Exchanger, Compressor and its suction KOD

& discharge cooler.

Expected CAPEX Lower Higher Highest

Expected OPEX Lowest Lower Highest

It may be noted that in all options, the stabilizer bottoms temperature is more than 100°C which can be used for heat integration with either Desalter Feed or Stabilizer Feed before cooled in downstream coolers (Air coolers and Trim coolers). The process inlet temperature to Air cooler is considered at 80°C and the calculated area of stabilizer/Desalter feed heat exchanger is compared for all cases.

The following are the inferences from the above table:

1. OPEX: The stabilizer reboiler duty in low pressure case (Option-3 @ 9.8 MW) is considerably lesser as compared to high pressure case (Option-1 @ 24.7 MW). However, this benefit is dwarfed by the advantage of high heat integration and reduced compressor duty in Option-1. Thus, total energy requirement (Hot Oil + Compressor duties) with higher operating pressure (Option-1 @ 32.7 MW) is lesser as compared to low operating pressure case (Option-3 @ 54 MW).

2. CAPEX: In Option-1, to enable high heat integration, heat exchanger area of 9500 m 2 will be required as compared to 2900 m2 in Option-3. However, this is nominal as compared to the additional Compressor, Suction K.O drum and Discharge cooler required in Option-3.





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In addition to the OPEX and CAPEX disadvantages indicated above, in Option-1, the calculated crude RVP (25.6 kPaa) is lesser than the allowable product specification (42.8 kPaa). This results in shrinkage of crude production for the same well fluid flow rate which is not acceptable.

Thus, the optimum stabilizer operating pressure is considered at 5.5 bara.

4.1.1.6 Feed Temperature to Stabilizer

The stabilizer feed stream coming from Desalter is at 116°C and the stabilizer bottom stream is available at 165°C. As a part of optimization of Stabilizer feed temperature and heat integration, the following two options are studied:

Option-1: Stabilizer feed heated to 133°C using Stabilizer Bottoms stream. The corresponding heat duty is 10 MW.

Option-2: Stabilizer feed at 116°C (same as Desalter outlet). Stabilizer bottom stream used to heat the partial draw-off from the stabilizer to an equivalent heat duty of 10 MW.

Refer to Annexure-3 for the Schematics representation of different Options.


Parameters Option-1 Option-2

Side Draw from - 4st Tray

Side Draw Return - 8th Tray

Stabilizer Feed Flow Rate, t/h 896 865

Stabilizer Feed Temperature, °C 133 120

Stabilizer Bottom Temperature, °C 165 165

Stabilizer Top Vapor Temperature, °C 124 113.5

Stabilizer Top Vapor Flow rate, m3/h 8028 5340

Stabilizer Top Vapor Molecular Weight 54.4 44

Compressor Duty, MW 8 7.9

Stabilizer Reboiler Duty, MW 24.7 20.7

Total Energy, MW 32.7 28.6

Feed Bottom Exchanger Duty, MW 10 -

Feed Bottom Exchanger Area, m2 1200 -

Side draw Exchanger Duty, MW - 10

Side draw Exchanger Area, m2 - 1500

Side Draw Accumulator Size (ID x L), m - 3 x 9

Side Draw Pump Capacity - 600 m3/h, 300 kW





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Stabilizer Size (ID x H), m 5.3 x 19 4.2 x 23

Product RVP, kPaa 44 44

Expected CAPEX Low High

Expected OPEX High Low

It may be noted that since stabilizer column does not have a condenser, the feed tray location is considered as the top tray so as to provide reflux stream for the top tray. Consequently, the flashed vapor from the feed escapes the Stabilizer from the top tray without contributing to the staged separation in the column. As a result, with increased feed temperature in Option-1, the amount of flashed vapor in the feed increases resulting in more vapor product exiting the top tray. This ineffective utilization of feed heat input (10 MW) is compensated by some additional reboiler duty resulting in an overall reboiler duty of 24.7 MW in Option-1. Also the compressor load is marginally high in Option-1 as the overhead vapor rate is higher as compared to Option-2.

In Option-2, flashed vapor in the feed is reduced as the Desalter outlet at 116°C @ 16 bara is directly fed into the Stabilizer without heating. Instead of heating the feed, the stabilizer bottoms is used to provide the same heat input of 10 MW to the column by heating the partial liquid drawn from stage-4 from the Stabilizer.

With Option-2, the required reboiler duty is 20.7 MW which is 16% lesser than that in Option-1 and also the vapor-liquid traffic within the column is made more uniform resulting in a reduced column diameter of 4.2 m with a nominal increase in the column height to accommodate the partial draw-off and return streams. Option-2 will require a side draw accumulator and a recirculating pump; however the aforesaid OPEX benefit will make this option favorable.

4.1.1.7 Heat Integration in Oil separation Unit

The hot and cold streams available in the Oil separation unit are tabulated in the below table:

Hot Streams Tin (°C) Tout (°C)Available Duty

(MW)

Stabilized crude 170 40 60.5

Desalter Effluent 116 40 4.5

Total - - 65.0

Cold Streams Tin (°C) Tout (°C)Required Duty

(MW)

Desalter Feed 56 120 32.3

Stabilizer Side Draw 112 143 10

Desalter Fresh Water Feed 45 - -

Stabilizer Reboiling 136 170 24

Total - - 66.3





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Basis:

Considering a typical minimum approach temperature of 20°C (between hot outlet and cold inlet), heat integration option is studied for the tabulated hot and cold streams.

1. With the above basis, the major heat source i.e., the stabilized crude is used to supply the total duty required by the Desalter Feed and the Stabilizer Side Draw stream (42.3 MW) and rest of hot stream duty (65.0 - 42.3 = 22.7 MW) in the low temperature range (85°C to 40°C) is rejected into the Cold utilities (Air cooler and trim coolers).

2. Heat exchange of 1.7 MW is considered between Desalter effluent and Desalter Fresh water feed up to the approach temperature limit and the rest of the duty (2.8 MW) in the cooling water exchanger.

3. The reboiler heat duty of 24.0 MW is supplied by the hot utility stream i.e., Circulating Hot Oil.

4.1.1.8 Gas Compression Unit

The following flashed vapor from the Oil Separation unit (1st Stage Separator at 15.5 bara and both 2nd

Stage Separator & Stabilizer overhead gas at 5.5 bara) is routed to the downstream GTP for producing Sales gas and LPG from associated gas. As the pressure requirement in the GTP, governed by the minimum pressure required in the De-ethaniser column (Refer section XX below), is 35 bara, the flashed gas is compressed in two-stages (pressure ratio per stage assumed as 3:1) to generate compressor discharge pressure of 45 bara to account for additional pressure drop in the units upstream of De-ethaniser column (35 bara).

The 2-stage compressor section is modeled with the following consideration:

1. The 2nd Stage Separator flashed gas and Stabilizer overhead gas at 5.5 bara are routed to the suction of 1st stage compression. Considering, a pressure ratio of 3:1, the calculated compressor condition is 16 bara @ 105°C (to be confirmed by Compressor Vendor).

2. The 1st stage compressor discharge is air cooled to 56°C to maintain the 2nd stage compressor discharge temperature within limit and to improve the compressor efficiency. The two phase stream at Intercooler outlet is flashed in the 2nd Stage compressor suction drum.

3. The flashed gas from 1nd Stage Separator and 2nd Stage compressor suction drum at about 15.5 bara is compressed in the 2nd stage compression to 45 bara @ 111°C (to be confirmed by Compressor Vendor).

4. The gas from 2nd stage compressor discharge (45 bara @111°C) should be routed to the downstream Gas Sweetening unit for H2S and CO2 removal. As low temperature favors the absorption in Amine Absorber tower, the compressor discharge is cooled to 40°C using Air cooler and Trim cooler in the compressor unit.

5. At 40°C, 2nd stage compressor discharge generates two phase which is routed to a 2nd Stage discharge drum for separation. The flashed sour gas (3.7 mol% H2S) is routed to the Gas Sweetening unit after superheating it by 10°C (using 2nd Stage Compressor discharge) to avoid H2S condensation in the interconnecting piping between compression unit and Gas Sweetening unit.





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6. The sour liquid (2.5 wt% H2S) from the 2nd Stage discharge drum (41 bara) is recycled back to the 2nd Stage suction drum (15 bara) and likewise, the total flashed liquid generated in the 2nd Stage suction drum is recycled back to the 1st Stage suction drum (5.0 bara). This two stage recycling of sour liquid from high pressure to low pressure enriches the Gas Sweetening unit feed gas with H2S and consequently reduces the H2S concentration in the liquid from 2.5 wt% observed in 2nd

Stage discharge drum to 0.2 wt% in the 1st Stage suction drum due to flashing. This liquid from 1st Stage suction drum (300 m3/hr.) forms the net sour liquid exiting the compressor section and the following consideration is used to decide its routing.

7. The sour liquid composition (70wt% C5+, 24wt% lighters, 0.2wt% H2S and 6wt% H2O) indicates that it is rich in C5+ components; hence most of this stream should be going with the stabilized crude. Accordingly, this liquid is routed to the 2nd Stage Separator inlet so that 6wt% H2O is removed sequentially in the Separator & Desalter and 24wt% lighters are stripped off in the Stabilizer.

Refer Appendix-3 for the schematics of the Gas Compression section.

4.1.2 Gas Treatment Plant

The net sour gas from the Associated Gas compressor with the composition as indicated in the table below forms the feed to Gas Treatment Plant (GTP). The objective of the GTP is to produce Sales gas and LPG from the associated gas meeting the product specification as mentioned in the BOD.

GTP Feed Stream composition:

Component Mol%

H2S 3.7

CO2 5.0

COS 0.0119Lighters 71.5LPG (C3s + C4s) 18.3C5+ 1.47Methyl Mercaptan 0.0135Ethyl Mercaptan 0.00261H2O 0.21Molecular Weight 26.97

GTP Product specification:

Sales Gas (as per BOD Table 8-4):





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Parameter Unit Specification

Water Dew Point @ 70 bar °C -12

Hydrocarbon Dew Point @ 70 bar °C -8

H2S ppmv <7.5

Mercaptan (RSH) ppmv <15

CO2 vol-% <2.5

Liquefied Petroleum Gas, LPG (as per BOD Table 8-5):

Component Unit Specification

Reid Vapor Pressure [RVP] - Summer kPaa 800

Reid Vapor Pressure [RVP] – Winter kPaa 1000

Ethane Vol-% <0.6

C5+ Vol-% <2.0

Sulphur mg/m3 <100

Water Content Vol-% 0 (Water Free)

4.1.2.1 Gas Sweetening and Solvent Regeneration Unit

The objective of the Gas Sweetening unit is to remove H2S and CO2 to meet the Sales Gas and LPG product specification. There are several processing technologies commercially available to treat the acid gas such as chemical absorption, physical absorption, physical adsorption and Permeation etc. While proprietary state of the art technology from different Licensors will be compared to select the final treatment facility, the simulation is carried out considering an open-heart chemical absorption based on Amine solvent in PROMAX software.

Feed Temperature:

As indicated in Gas compression section above, the Gas Sweetening unit feed is supplied at 50°C





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The flashed vapor from the





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4.2 Gas Sweetening Unit

4.3 Gas Dehyration Unit

4.4 HC Dew Point Control Unit

4.5 Deethaniser Column

4.6 Depropaniser Column

4.7 Debutaniser Column

4.8 Propane Regfrigeration Unit

4.9 Sensitivity Study with PVT Data





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process stimulation

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